Process for the production of hydrogen and the co-production of carbon dioxide

ABSTRACT

Process for the production of hydrogen and the co-production of carbon dioxide in separate streams comprising the following operations: 
         feeding of a solid to a first reaction zone (R 1 ) in which a liquid or gaseous hydrocarbon, preferably natural gas or methane, is also fed, which reacts with said solid fed at its maximum oxidation degree (over-reduced form), leading to the formation of the combustion products carbon dioxide and water and the solid at its minimum oxidation degree (reduced form);    feeding of the solid in reduced form to a second reaction zone (R 2 ) into which water is also fed, which reacts with said reduced form of the solid, producing hydrogen, steam and the solid at an intermediate oxidation degree (oxidized form);    feeding of the solid in oxidized form to a third reaction zone (R 3 ) into which air is also fed, obtaining, from the further oxidation of the solid, heat and the solid in over-oxidized form to be recycled to the first reaction zone (R 1 ), wherein said solid contains at least on element selected from elements which, in addition to the metallic state, have at least three different oxidation states and are therefore capable of producing two redox pairs, in order of the oxidation state, 
 
characterized in that in the first reaction zone (R 1 ) the solid in over-oxidized form and the liquid or gaseous hydrocarbon are fed in countercurrent.

The present invention relates to a process which uses natural gas ormethane and water for the production of hydrogen and the co-productionof carbon dioxide in separate streams.

The H₂ produced according to the invention can be used not only in allhydrogenation operations but also for feeding energy generators such asfuel cells, whereas the CO₂ co-produced can be eliminated where mostconvenient (exhausted well, sea bottom, etc.) to avoid its accumulationin the atmosphere.

The known art describes processes and technologies for the production ofsynthesis gas CO/H₂ by the direct oxidation of natural gas using steam(steam reforming) or oxygen (partial oxidation) as oxidizing agent.

In the first case (steam reforming)CH₄+H₂O═CO+3H₂the thermodynamic limits established by the formation of equilibrium,must be faced in the reaction areaCO+H₂O←→CO₂+H₂In the second case (partial oxidation)CH₄+1/2O₂=CO+2H₂problems of selectivity must be faced with respect to the totaloxidation of the natural gas to CO₂/H₂O which can only be overcome byoperating within a temperature range (over 1000° C.), which, in additionto the necessity of a cryogenic air fractionation unit, makes theindustrial technological application onerous.

In both cases, downstream of the natural gas conversion, the CO/H₂mixture must still be treated with steam to increase the production ofH₂ (low/high temperature WGS) followed by a first physical separation(PSA) of the hydrogen and subsequently, to obtain a high purity product,a further separation of the residual CO₂ by chemical absorption (washingwith amines). Furthermore, in the case of steam reforming, a thermalsupport must be provided for the process as the reforming step anddesorption from the amine of the carbonated CO₂ are of an endothermicnature.

In order to overcome problems relating to thermal support, autothermalreforming processes are being widely adopted, which envisage the feedingof the CH₄/O₂ mixture in a molar ratio (2:1) but operating within atemperature range which is such as to favour a total combustion reactionCH₄+2O₂=CO₂+2H₂Otogether with which the reforming of the excess methane with the H₂O andCO₂ produced by the combustion, is effected contemporaneously.

In this way, the exothermy of total oxidation is balanced with theendothermy of steam reforming; in this case, with respect to partialcombustion, there is no technological problem caused by the combustiontemperature but a cryogenic air fractionation unit is still necessary.

The known art also describes processes and technologies for theproduction of H₂/CO₂ by the indirect steam reforming of natural gas;these are characterized by the formation of products in separatestreams, thus not only avoiding all separation and purificationoperations of the products (HT/LT WGS; PSA; amines) but also overcomingall the thermodynamic limits established by the WGS reaction which runsconsecutively with the direct reforming reaction discussed above.

In these processes, the reaction between H₂O and natural gas is indirectas the oxygen exchange takes place through a solid capable of providingone or more intermediate redox pairs (Meo/Me_(x)O_((x+1)); Me/MeO)between the reducing potential of the CH₄/CO₂ pair and the oxidizingpotential of the H₂O/H₂ pair. In practice, the redox solid acts asoxygen donor to the natural gas (reducing agent) and oxygen receiverfrom the water (oxidizing agent). The oxygen exchanged through the solidis chemically defined as “reversible oxygen” hereafter indicated asC.O.A. (chemical oxygen available) when it is released and as C.O.D.(chemical oxygen demand) when it is acquired by the solid.

Patent U.S. Pat. No. 3,442,620 describes a process in which, in a firstphase (endothermic) which leads to the production of the CO₂/H₂O stream,the redox solid in its oxidized form, is completely or partially reducedto the metallic state by a reducing agent (in this case synthesis gasproduced from coal and consequently there are two redox pairs inquestion (CO/CO₂ and H₂/H₂O).

The main reduction scheme is described by the following equations:Me_(x)O_((x+1))+CO ←→X MeO+CO₂  a)Me_(x)O_((x+1))+H₂←→X MeO+H₂O  b)MeO+CO←→Me+CO₂  c)MeO+H₂←→Me+H₂O  d)In a second process phase (exothermic) which leads to the production ofthe H₂ stream, the redox solid in its reduced form is oxidized by theH₂O to the initial oxidation state according to the following equations:Me+H₂O←→MeO+H₂  e)XMeO+H₂O←→Me_(x)O_((x+1))+H₂  f)wherein Me refers to Fe and X=3.

In the first process step, the solid oxygen carrier (oxidizing agent) inits most oxidized form (Me_(x)O_((x+1)) in down flow) and the reducingagent (CO/H₂ in up flow) are fed in countercurrent to each other in asub-stoichiometric Ox/Red molar ratio with each other (reducing defect).

The reduction of the solid which is carried out with a step-wisemechanism, takes place in a two-zone reactor, a first upper zone wherethe formation of MeO (reactions a and b) is effected and a second lowerzone where the further reduction to Me is effected (reactions c and d).

In the second process step, the solid oxygen receiver (reducing agent)obtained in the first step (Me/MeO mixture in down flow) and theoxidizing agent (H₂O in up flow) are fed in countercurrent to each otherin an overstoichiometric Ox/Red molar ratio with each other (oxidizingexcess) with the partial transformation of the oxidizing agent to H₂.

The oxidation of the solid is carried out in two zones, a first upperzone where the formation of MeO (reaction e) takes place and a secondlower zone where the further oxidation of the solid is effected to itsinitial form (reaction f).

These process zones however cannot be considered in the light ofsubsequent process steps as each zone is characterized by a set ofreactions at equilibrium.

In addition to these reactive steps, the process comprises a riser fedwith an inert or poorly reactive gas such as the reduction spent-gas (amixture of CO₂/H₂O/CO/H₂) which allows the solid obtained in the secondstep (simple pneumatic conveyance) to be re-fed to the first processstep thus obtaining the continuous production of the two gaseous processstreams described above.

All the process steps (a-f) are in equilibrium; consequently, as theprocess operates at “autogenous” conversion values (of equilibrium),once the pressure has been set within a range of 200-2500 psig, in orderto close the process balance, a pair of temperatures must be selected(also on the basis of the composition of the reducing gas —H₂/CO ratio)within a range of 1000° F.-2000° F. for carrying out the reduction stepand oxidation step respectively.

As the process operates at a high solid/gas ratio and with a partialconversion of the reducing gas, as a consequence:

-   -   the reduction off-gas does not consist of an intrinsically        removable stream of CO₂    -   the conversion per passage of exchangeable oxygen (redox yield)        is rather low (values around 10-15%) which means a low        yield/time-space to hydrogen.

In conclusion the chemical process yield is in the order of 65%.

Furthermore the problem of the closing of the thermal balance remainsunsolved as the overall process is endothermic (the process reaction isCO+H₂O═CO₂+H₂).

Patent application EP-1134187 also proposes the indirect reactionbetween H₂O and natural gas and describes a process and material whereinin a first step (endothermic) which leads to the production of thestream of CO₂ and H₂O, the redox solid in its oxidized form is reducedin counter-current by a reducing agent in upflow (in this case CH₄ andconsequently there is the redox pair CH₄/CO₂).

The main reduction scheme is described by the following equation:4Me_(x)Z_(z)O_((y+1))+CH₄←→4Me_(x)Z_(z)O_(y)+CO₂+2H₂O

-   -   wherein x≧1; y≧0; z≧0    -   wherein Me represents the redox element which is preferably        supported Fe    -   wherein Z acts as promoter preferably selected from Ce, Cr, Zr.

More specifically, the reduction step of the material contemplates, withrespect to the oxidizing agent, four parallel equilibriums (h; I; j; k),as follows:4Fe₃O₄+CH₄←→12 FeO+CO₂+2H₂O Kh=[CO₂][H₂O]/[CH₄]  h)Fe₃O₄+CH₄←→3 FeO+CO+2H₂ Ki=[CO][H₂]²/[CH₄]  i)Fe₃O₄+CO←→3 FeO+CO₂Kj=[CO₂]/[CO]  j)Fe₃O₄+H₂←→3 FeO+H₂O Kk=[H₂O]/[H₂],  k)which, on the other hand, are of the competitive type with respect tothe reducing species giving rise to possible problems of sub-conversionfor the kinetically unfavourable species (rate determining steps).

To complete the conversion of the reducing agent, the technologyproposed optionally envisages the introduction of an enrichmentoperation (TSA; PSA) of the reduction spent-gas in the reducing species(H₂/CO) and to subsequently re-feed it with further fresh redox solid(in a substantial excess) to an additional reduction zone. This howeveris onerous from an economical point of view.

In a second step of the process (exothermic) which leads to theproduction of the stream of H₂, the reduced redox solid is oxidized tothe initial oxidation state in countercurrent with steam in upflowaccording to the following equation:4Me_(x)Z_(z)O_(y)+4H₂O←→4Me_(x)Z_(z)O_((y+1))+4H₂  1)

With respect to U.S. Pat. No. 3,442,620, a particular selection of thepair of reduction and oxidation temperatures for closing the molarbalance, is not required as the reduction is effected at completeconversion (not a spontaneous equilibrium conversion) of the reversibleoxygen and in the oxidation, the conversion of the steam can beconsidered as a freedom degree.

As the overall process is endothermic (the reaction is 2H₂O+CH₄→CO₂+4H₂)the patent proposes the closing of the thermal balance by interposing,prior to the two redox steps, a thermal supporting riser (third step) inwhich heat is supplied to the process for:

-   -   combustion of part of the hydrogen produced    -   combustion of additional natural gas    -   over-oxidation with air of the redox solid    -   feeding an air combustor with reduction spent-gas containing        CO/H₂ not before, however, effecting the CO₂ and H₂O operation        (PSA etc.)

In conclusion, the chemical process yield is around 83% (enthalpicefficiency close to 90%).

Although the technology described allows the objectives of said Europeanpatent application to be reached with extremely high chemical yield andenthalpic values, the problem remains however of thecontrol/minimization (at the stoichiometric value) of the MO/CH₄ ratio,as may be required for the reduction reaction on various cascade zonesand operating in excess of reversible oxygen whose direct consequence isa decrease in the space/time yield on the redox solid due to the lowconversion of the oxidized form Me_(x)Z_(z)O_((y+1)) fed to thereduction step with methane (reaction g)).

Patent application IT-MI03A000192 describes the preparation of reactivesystems consisting of an active phase based on Iron and amicrospheroidal carrier based on micro-spheroidal alumina capable ofproviding a high redox yield (productivity to hydrogen in terms ofNltH₂/Kg of material) which are processed with a three-reaction-zoneprocess comprising a first process phase in which water (oxidizing agentand solid (reducing agent) enter and H₂ is produced together withoxidized solid; a second process phase for heat supply in which air(oxidizing agent) and oxidized solid (reducing agent) enter andproducing heat together with over-oxidized solid; a third process phasein which the over-oxidized solid (oxidizing agent) enters with naturalgas (reducing agent) producing CO₂/H₂O and reduced solid.

The high productive capacity of these materials can only be obtained,however, if these are processed with suitable reactor expedients bothfrom a fluid-dynamic (promoting the fluidization of the reactive bed andtherefore promoting the oxygen exchange between gas and solid) andchemical point of view, as the use of reactive systems based on ironoxides contemplates the presence of parallel equilibrium reactions (h;I; j; k) which, in the absence of expedients to shift the equilibriumstowards the formation of products, cause a low process yield.

We have found an advantageous process which can be applied on anindustrial scale and which allows the continuous production of separatestreams of H₂ and CO₂ with a high purity, at the same time, maximizingthe space and time yield, as there is no necessity of resorting eitherto separation or purification operations downstream.

The process, object of the present invention, for the production ofhydrogen and the co-production of carbon dioxide comprises the followingoperations:

-   -   feeding of a solid to a first reaction zone (R1) in which a        liquid or gaseous hydrocarbon is also fed, which reacts with        said solid fed at its maximum oxidation degree (over-reduced        form), leading to the formation of the combustion products        carbon dioxide and water and the solid at its minimum oxidation        degree (reduced form);    -   feeding of the solid in reduced form to a second reaction zone        (R2) into which water is also fed, which reacts with said        reduced form of the solid, producing hydrogen, steam and the        solid at an intermediate oxidation degree (oxidized form);    -   feeding of the solid in oxidized form to a third reaction zone        (R3) into which air is also fed, obtaining, from the further        oxidation of the solid, heat and the solid in over-oxidized form        to be recycled to the first reaction zone (R1),        wherein said solid contains at least one element selected from        elements which, in addition to the metallic state, have at least        three different oxidation states and are therefore capable of        producing at least two redox pairs, according to the oxidation        state,        characterized in that in the first reaction zone (R1) the solid        in over-oxidized form and the liquid or gaseous hydrocarbon are        fed in countercurrent.

The hydrocarbon fed in the process according to the invention ispreferably gaseous, more preferably natural gas or methane.

The redox reaction of the first reaction zone (R1) can be carried out ina step reactor, preferably at pressures ranging from 1 to 20 bar, attemperatures lower than or equal to 900° C., continuously feeding incountercurrent, over-oxidized solid (C.O.A)/natural gas or methane witha molar ratio less than or equal to 4/1.

The concept of a step reactor envisages the fact that a single reactionis effected on a series of successive chemical equilibrium steps at anincreasing conversion degree.

The water and carbon dioxide can be removed from the top of the stepreactor with a molar ratio CO₂/H₂O equal to 1/2.

The redox reaction of the second reaction zone (R2) can be carried outin a step reactor, preferably operating at a pressure substantiallyequal to that of the step reactor of the first reaction zone, in whichthe reduced solid (C.O.D.) and the steam are fed countercurrent incontinuous with a molar ratio less than or equal to 1.

The redox reaction of the third reaction zone (R3) can be carried out ina riser, feeding air and the oxidized solid in equicurrent.

The additional heat supply to the system can also be obtained by thecombustion of natural gas or methane optionally co-fed with air into thethird reaction zone (R3).

Solids which can be used are those containing at least one elementselected from elements having at least three different oxidation states,stable under the reaction conditions, which differ in their oxygencontent and in that they are capable of cyclically passing from the mostreduced form to the most oxidized form and vice versa.

Solids containing one or more elements with the above characteristicscan be used, i.e. having, in addition to the metallic state, at leastthree different oxidation states, preferably three states, and capableof producing in the order of oxidation state, at least two redox pairs,preferably 2 pairs, and can be adopted as such or in a mixture withother elements which are not subject to redox reactions; the reactivephase thus obtained can, in turn, be used as such or suitably dispersedor supported on compounds such as silica, alumina, or other pure oxidessuch as magnesium, calcium, cerium, zirconium, titanium, lanthanum, butalso mixtures thereof.

Among solids having at least three different oxidation states, ironproves to be particularly advantageous, and can be present in the solidin binary formFe_(x)O_(y)And/or in ternary formFe_(x)Z_(z)O_(y)Wherein x≧1, y≧0, z≧1,Z is at least an element selected from Ce, Zr, V and Mo.

In the third reaction zone (R3), the element selected from elementshaving at least three different oxidation states can optionally consistof two phases deriving from the fact that the oxidation step in R2 iscarried out with a partial and incomplete conversion of the element:when the element is iron, the two phases are FeO and Fe₃O₄.

A preferred embodiment according to the invention can be obtained by aprocess configuration wherein the first phase that leads to thereduction of the redox solid (reaction (g)) is carried out in a stepreactor which operates under suitable conditions of P (from 1 to 20 bar)and temperature (up to a maximum of 900° C. and depending however on theactive phase used) to which natural gas or methane (reducing agent) arefed in continuous in upflow, and the reversible oxygen carrierMe_(x)Z_(z)O_((y+1)) (solid oxidizing agent) in downflow in a molarratio (RED/OX) between each other that can be lower than or equal tothat given by the stoichiometry of the reaction (g) (1/4), and a streamof CO₂/H₂O in a molar ratio equal to 1/2, is continuously removed, fromabove and, a stream of reduced solid Me_(x)Z_(z)O_(y), from below.

In this way, the reduction reactions (h; I; j; k) of Fe₃O₄ with CH₄which lead to the formation of the hydrogen precursor (FeO) are effectedon a series of equilibrium steps at an increasing conversion degree(reduction) in downflow for the oxidizing solid and increasingparameters (oxidation) in upflow for the reducing gas.

From a chemical point of view, it is verified that in upflow, thereducing potential of the gas decreases due to the formation of theredox pairs CO₂/CO and H₂O/H₂, but, as the pair Fe₃O₄/FeO increases, theoxidizing potential of the solid, which actually supports the reaction(greater availability of oxygen which moves the equilibrium to theright), increases.

In order to provide thermal support to the process with the introductionof a third over-oxidation phase of the reactive system, in the upperpart of the reactor (upflow), there is an additional and even moreefficient redox pair Fe₂O₃/Fe₃O₄, capable of producing and sustainingthe reaction even operating at high gas/solid ratios.

From a reactor point of view, the situation can be effected in a reactorof the plug-flow type which typically operates at a conversion profilewhich varies along the in/out axis of the reagent feeding and whoseefficiency is linked to the minimization of back-mixing phenomenabetween solid and gas. In principle, a mobile bed reactor orfractionated (staged) fluid bed reactor can be used by the introductionof diaphragms (for example perforated plates or others), whose pressuredrop (passage span) is studied to limit back-mixing phenomena and at thesame time, allowing the upward movement of the gas and downward movementof the solid. The distance between one diaphragm and another (holdup ofeach step) is calculated on the basis of the kinetic characteristics ofthe reaction so as to obtain, for each step, an optimum distribution ofthe overall contact time. The minimum number of steps required is thatfixed by the thermodynamics of the reactive system at the variousreaction temperatures which are established in the reduction reactor andare determined by the thermal balance on said reactor.

The second phase of the process which leads to the oxidation of theredox solid (reaction (k)), is carried out in a multi-step reactor whichoperates at the same reactor pressure as the previous phase, and at atemperature selected on the basis of the thermodynamics and kinetics ofthe reaction (k), to which steam (oxidizing agent) is fed in continuousin upflow and the solid oxygen receiver Me_(x)Z_(z)O_(y) (reducingagent) in downflow, in a molar ratio (RED/OX) which can be lower than orequal to, preferably lower than 1 (excess oxidizing agent), at thestoichiometry of the reaction (k), and a stream of H₂ and steam isremoved in continuous from above, whereas a stream ofMe_(x)Z_(z)O^((y+1)) is removed from below, which can be re-fed directlyto the previous phase.

From the point of view of fluid-dynamics, this phase of the process canalso be carried out in a fluid bed reactor which, according to the art,can be equipped with dividers whose function is to force thefluidization of the solid and thus improve the oxygen exchange betweensolid and gas (reduction in the diameter of the gas bubbles) There canbe various types of dividers (for example, perforated plates; chevron)depending on the rheological characteristics of the gas and solid.

The third phase, where, in order to close the thermal balance of theprocess, the over-oxidation takes place of the active phase to Fe₂O₃with air (exothermic), can be carried out by feeding gas and solid inequicurrent to a riser with the subsequent re-entry of this solid to thefirst phase of the process, preferably by pneumatic conveyance.

An embodiment of the present invention is provided below with the helpof FIG. 1, which should in no way be considered as limiting the scope ofthe invention itself.

With reference to the loop illustrated in FIG. 1, R1 represents thefirst process phase for the reduction of the solid and production ofCO₂, R2 represents the second process phase for the oxidation of thesolid and production of H₂, R3 represents the third process phase forthe thermal support by over-oxidation of the solid.

The operating pressure of the loop is 20 ata.

Methane (1) is fed to the first phase of the process (R1), from whichthe combustion products CO₂ and H₂O (2) are removed; steam (3) is fed tothe second phase of the process (R2), from which the oxidation productH₂ (4) is removed; air (5) is fed to the third phase of the process(R3), from which impoverished air (6) is removed.

The scheme is completed by the circulation lines of the solid whichconnect the three process phases, the reduced solid (8) coming from thereduction phase (R1), is fed to the oxidation phase, the oxidized solid(9) coming from the oxidation phase (R2), is fed to the over-oxidationphase, the over-oxidized solid (7) coming from the heat production phase(R3), is fed to the reduction.

EXAMPLE 1

With reference to the scheme of FIG. 1, the operating pressure of theloop is 20 ata, methane (1), steam (3) and air (5) are fed afterpreheating. The circulation flowrate of the solid is in relation to theFe₂O₃/CH₄ ratio necessary for completely converting the methane to CO₂and water in the upper part of R1 (ratio depending on both thethermodynamics and contact time in R1) but at the same time it must besuch as to guarantee the closing of the process thermal balance withoutresorting to excessively high ΔT values in the endothermic step(reduction) (Minimum temperature in R1—lower part—not lower than 710°C.) to avoid problems relating to kinetics and also excessively high ΔTvalues in the exothermic step (over-oxidation)) (Maximum temperature inR3—upper part—not exceeding 850° C.). In accordance with this, redoxsolids are used, characterized in their thermal capacity values (Cp) andin their active phase (Fe₂O₃)/inert carrier ratio, capable therefore ofreleasing or acquiring (excluding the formation of metallic phases) aquantity of oxygen (O reversible) equal to a few weight percentage units(>1% w) with respect to the total of the carrier (active phase+inertcarrier).

From the point of view of chemical-performance, the productivity of thisredox solid is equal to about 14 NLt of H₂ per Kg of solid processed perpercentage of reversible oxygen acquired in the oxidation step withwater (C.O.D.).

For R1, a multi-step reactor is used in countercurrent, in which theminimum number of steps (theoretical steps) is in relation to thepositioning of the chemical equilibriums at the operating temperaturesin which the stable reduced species is FeO, passing through theformation of the sub-stoichiometric oxide Fe_(0.947)O (see followingtables); for R2, on the other hand, a single equilibrium step has alwaysbeen considered: in this reactor, in fact, by stopping the reduction ofFe₃O₄ in R1 at FeO, there is only the inverse reaction from FeO to Fe₃O₄and there is therefore no advantage in using a multi-step reactor (theremay be this advantage however from a fluid-dynamic point of view).

The air flow-rate to R3 is selected case by case on the basis offluid-dynamic considerations to avoid having an excessively denseconveyance line. TABLE 1 Reduction phase in R1 and production of CO₂feeding 2379 kg/hr of REDOX solid and 22.4 Nmc/hr of methane (R 4.47:1)Solid Inlet gas Step 4 inlet line (1) Step 1 Step 2 Step 3 line (4) line(7) T° C. 450 724.7 783.4 839.0 853.7 848 Flow rate kmoles/hour CH₄ 10.480 0.121 0.018 0.0 CO₂ 0 0.268 0.487 0.582 1.0 H₂O 0 0.444 0.9201.199 2.0 CO 0 0.252 0.391 0.400 0.0 H₂ 0 0.596 0.838 0.766 0.0 Fe₂O₃ 00 0 0.761 4.47 Fe₃O₄ 0.47 1.444 2.710 2.472 0 FeO 7.53 0 0 0 0Fe_(0.947)O 0 4.867 0.854 Inert carrier 11.708 11.708 11.708 11.70811.708

TABLE II Oxidation phase in R2 and production of H₂ by feeding 2314kg/hr of REDOX solid reduced and 83.34 kg/hr of H₂O. Inlet gas Line (3)Step 1 Line (4) Solid inlet Line (8) T° C. 400 734.8 707.6 Flow ratekmoles/hour H₂O 4.63 2.12 H₂ 0 2.51 Fe₂O₃ 0 0 Fe₃O₄ 2.98 0.47 FeO 0 7.53Inert carrier 11.708 11.708

TABLE III Over-oxidation phase of the REDOX solid in R3 and productionof heat by feeding 2354.9 kg/hr of oxidized REDOX solid and 136 Nmc/hrof air. Inlet gas Inlet solid Out gas Out solid line (5) line (9) line(6) line (7) T° C. 430 734.8 848 848 Flow rates kmoles/hour N₂ 4.8194.819 O₂ 1.281 0.536 Fe₂O₃ 0 4.47 Fe₃O₄ 2.98 0 FeO 0 0 Inert carrier11.708 11.708

From the values of Tables I, II and III it can be observed that:

-   -   The methane conversion at the end of the reactive steps with        Fe₃O₄ (steps 1÷3) is 98.2% whereas the selectivity to CO₂ and        H₂O is equal to 58%;    -   The methane conversion at the end of the reactive step with        Fe₂O₃ (step 4) is 100% as also the selectivity to CO₂ is 100%;    -   The conversion of the water in R2 is 54.2%;    -   The oxygen released (C.O.A.) from the solid in R1 (Δp) is −2,7%        (90% of the theoretical exchangeable oxygen value in the        Fe₂O₃—FeO passage); that received (C.O.D.) from the solid in R2        (Δp) is +1.7% and that received in R3 (Δp) is, by difference,        +1.0%;    -   2.51 Lt of H₂ are produced per litre of methane processed and        consequently the thermal efficiency of the cycle, referring to        (QH₂xΔH_(c)H₂)/(QCH₄xΔH_(c)CH₄) is 75.65%;    -   23.63 NLt of H₂ are produced per Kg of solid processed and        consequently the productivity of the solid is 84.4%    -   The discharge of the gases from the head of R1 (line 2) consists        of a stream of which a third consists of CO₂ and two thirds of        H₂O: after cooling and condensation, this stream forms a stream        of pure CO₂;    -   The discharge of the gases from R2 (line 4) consists of an        almost equimolar stream in steam and hydrogen: also in this        case, after cooling and condensation, a stream is obtained which        is practically pure in H₂.

The example shows that, when operating with a temperature of 853.7° C.in the upper part of R1, 4 equilibrium steps are sufficient. In order tocompletely convert the methane with a selectivity to CO₂ and H₂O equalto 100%, it is necessary however to operate with respect to thestoichiometric value, in excess of Fe₂O₃ (4.47:1).

EXAMPLE 2

Reference is made to the scheme of FIG. 1 and with the same assumptionsmade in Example 1 except for the fact that in this case R3 represents aover-oxidation unit with air of the REDOX solid for the thermal supportof the process, but, unlike the previous case, methane is also fed asfuel (mixed thermal support). TABLE IV Reduction phase in R1 andproduction of CO₂ by feeding 2132 kg/hr of REDOX solid and 22.4 Nmc/hrof methane (R 4.0:1). Inlet gas Step 5 Solid inlet line (1) Step 1 Step2 Step 3 Step 4 line (2) line (7) T° C. 450 707.5 750.7 799.7 843.5854.7 848 Flow rate kmoles/hour CH₄ 1 0.599 0.308 0.070 0.015 0 CO₂ 00.203 0.364 0.527 0.587 1 H₂O 0 0.319 0.641 1.024 1.215 2 CO 0 0.1970.328 0.403 0.398 0 H₂ 0 0.482 0.743 0.835 0.755 0 Fe₂O₃ 0 0 0 0 0.359 4Fe₃O₄ 0 0.572 1.500 2.444 2.427 0 FeO 8 0 0 0 0 0 Fe_(0.947)O 0 6.6373.696 0.705 Inert carrier 10.477 10.477 10.477 10.477 10.477 10.477

TABLE V Oxidation phase in R2 and production of H₂ by feeding 2110.9kg/hr of reduced REDOX solid and 86.22 kg/hr of H₂O. Inlet gas Line (3)Step 1 Line (4) Solid inlet Line (8) T° C. 400 734.0 707.5 Flow ratekmoles/hour H₂O 4.79 2.123 H₂ 0 2.667 Fe₂O₃ 0 0 Fe₃O₄ 2.667 0 FeO 0 8Inert carrier 10.477 10.477

TABLE VI Over-oxidation phase of REDOX solid in R3 and production ofheat by feeding 2354.9 kg/hr of oxidized REDOX solid and 121.85 Nmc/hrof air and 0.87 Nmc/hr of methane. Inlet gas Fuel Solid inlet Out gasOut solid line (5) line (5) line (9) line (6) line (7) T° C. 430 450 734848 848 Flow rate kmoles/hour N₂ 4.298 4.298 O₂ 1.142 0.398 CH₄ 0.039 0CO₂ 0.039 H₂O 0.078 Fe₂O₃ 0 4 Fe₃O₄ 2667 0 FeO 0 0 Inert carrier 10.47710.477

From the values of Tables IV, V and VI it can be observed that:

-   -   The methane conversion at the end of the reactive steps with        Fe₃O₄ (steps 1÷4) is 100% whereas the selectivity to CO₂ and H₂O        is equal to 58%;    -   The methane conversion at the end of the reactive step with        Fe₂O₃ (step 5) is 100% as also the selectivity to CO₂ is 100%;    -   The conversion of the water in R2 is 55.7%;    -   The oxygen released (C.O.A.) from the solid in R1 (Δp) is −3,0%        (100% of the theoretical exchangeable oxygen value in the        Fe₂O₃—FeO passage); that received (C.O.D.) from the solid in R2        (Δp) is +2.0% and that received in R3 (Δp) is, by difference,        +1.0%;    -   2.56 Lt of H₂ are produced per litre of methane processed and        consequently the thermal efficiency of the cycle, referring to        (QH₂xΔH_(c)H₂)/(QCH₄xΔH_(c)CH₄) is 76.51%;    -   28.1 NLt of H₂ are produced per Kg of solid processed and        consequently the productivity of the solid is 100%    -   The discharge of the gases from the head of R1 consists of a        stream of which a third consists of CO₂ and two thirds of H₂O:        after cooling and condensation, this stream forms a stream of        pure CO₂;    -   The stream at the outlet of R3 also contains CO₂ and water, due        to the combustion of the CH₄ in R3: the separation of CO₂ is        consequently not total but is 96.3%;    -   The discharge of the gases from R2 (line 4) consists of an        almost equimolar stream in steam and hydrogen: also in this        case, after cooling and condensation, a stream is obtained which        is practically pure in H₂.

This example shows that it is possible, with a temperature of 854.7° C.in the upper part of R1, to completely convert the methane with aselectivity to CO₂ and H₂O equal to 100% and at the same time to obtainthe maximum productivity in H₂ of the solid, by feeding Fe₂O₃ at astoichiometric value with methane (4:1).

To obtain this, it is necessary to co-feed a quota of methane to thethermal support unit which envisages the presence of CO₂ (0.8%) in thefumes from R3.

Five equilibrium steps are necessary for this case (in practice theintroduction of an additional step with respect to the previous caselimits the excess of Fe₂O₃, from the point of view of the methane, theH₂/CH₄ productivity proves to paradoxically improve even if a quota isfed to R3).

EXAMPLE 3

Reference is made to the scheme of FIG. 1 and with the same assumptionsmade in Example 1 except for the fact that in this case R3 represents aover-oxidation unit for thermal support with air of the REDOX solidwhich consists of two phases (FeO and Fe₃O₄) deriving from the fact thatthe oxidation step is carried out in R2 with a partial and incompleteconversion of the FeO. TABLE VII Reduction phase in R1 and production ofCO₂ by feeding 2132 kg/hr of solid REDOX and 22.4 Nmc/hr of methane (R4.0:1). Inlet gas Step 5 Solid inlet line (1) Step 1 Step 2 Step 3 Step4 line (2) line (7) T° C. 450 707.6 751.7 802.9 847 856.8 850 Flow ratekmoles/hour CH₄ 1 0.604 0.301 0.063 0.013 0 CO₂ 0 0.197 0.368 0.5340.590 1 H₂O 0 0.309 0.650 1.043 1.226 2 CO 0 0.199 0.331 0.404 0.397 0H₂ 0 0.483 0.748 0.832 0.747 0 Fe₂O₃ 0 0 0 0 0.407 4 Fe₃O₄ 0 0.546 1.5262.483 2.396 0 FeO 8 0 0 0 0 0 Fe_(0.947)O 0 6.719 3.614 0.583 0 0 Inertcarrier 10.477 10.477 10.477 10.477 10.477 10.477

TABLE VIII Oxidation phase in R2 and production of H₂ by feeding 2110.9kg/hr of reduced REDOX solid and 83.08 kg/hr of H₂O. Inlet gas Line (3)Step 1 Line (4) Solid inlet Line (8) T° C. 400 734.8 707.6 Flow ratekmoles/hour H₂O 4.565 2.013 H₂ 0 2.552 Fe₂O₃ 0 0 Fe₃O₄ 2.552 0 FeO 0.3438 Fe_(0.974)O 0 0 Inert carrier 10.477 10.477

TABLE IX Over-oxidation phase of the REDOX solid in R3 and production ofheat by feeding 2354.9 kg/hr of oxidized solid REDOX and 123.6 Nmc/hr ofair. Inlet gas Inlet solid Out gas Out solid line (5) line (9) line (6)line (7) T° C. 430 743.2 850 850 Flow rates kmoles/hour N₂ 4.361 4.361O₂ 1.159 0.435 Fe₂O₃ 0 4 Fe₃O₄ 2.550 0 FeO 0.351 0 Fe_(0.974)O 0 0 Inertcarrier 10.477 10.477

From the values of Tables VII, VIII and IX, it can be observed that:

-   -   The methane conversion at the end of the reactive steps with        Fe₃O₄ (steps 1÷4) is 100% whereas the selectivity to CO₂ and H₂O        is equal to 58%;    -   The methane conversion at the end of the reactive step with        Fe₂O₃ (step 5) is 100% as the selectivity to CO₂ is also 100%;    -   The conversion of the water in R2 is 55.9%;    -   The oxygen released (C.O.A.) from the solid in R1 (Δp) is −3,0%        (100% of the theoretical exchangeable oxygen value in the        Fe₂O₃—FeO passage); that received (C.O.D.) from the solid in R2        (Δp) is +1.9% and that received in R3 (Δp) is, by difference,        +1.1%;    -   2.55 Lt of H₂ are produced per litre of methane processed and        consequently the thermal efficiency of the cycle, referring to        (QH₂xΔH_(c)H₂)/(QCH₄xΔH_(c)CH₄) is 76.93%;    -   26.9 NLt of H₂ are produced per Kg of solid processed and        consequently the productivity of the solid is 95.7%    -   The discharge of the gases from the head of R1 consists of a        stream of which a third consists of CO₂ and two thirds of H₂O:        after cooling and condensation, this stream forms a stream of        pure CO₂;    -   There is the total separation of CO₂;    -   The discharge of the gases from R2 consists of an almost        equimolar stream in steam and hydrogen: also in this case, after        cooling and condensation, a stream is obtained which is        practically pure in H₂.

This example shows that it is possible to completely convert the methanewith a selectivity to CO₂ and H₂O equal to 100% by feeding Fe₂O₃(C.O.A.) in a stoichiometric ratio with methane (4:1).

In order to obtain this, the maximum production of H₂ in R2 must berenounced and consequently not only Fe₃O₄ but also residual FeO willenter R3. This will allow the total thermal balance to be closed withoutthe necessity of thermal support.

Also in this case, the minimum number of equilibrium steps necessaryproved to be equal to 5.

1. A process for the production of hydrogen and the coproduction of carbon dioxide comprising the following operations: feeding of a solid to a first reaction zone (R1) in which a liquid or gaseous hydrocarbon is also fed, which reacts with said solid fed at its maximum oxidation degree (over-reduced form), leading to the formation of the combustion products carbon dioxide and water and the solid at its minimum oxidation degree (reduced form); feeding of the solid in reduced form to a second reaction zone (R2) into which water is also fed, which reacts with said reduced form of the solid, producing hydrogen, steam and the solid at an intermediate oxidation degree (oxidized form); feeding of the solid in oxidized form to a third reaction zone (R3) into which air is also fed, obtaining, from the further oxidation of the solid, heat and the solid in over-oxidized form to be recycled to the first reaction zone (R1), wherein said solid contains at least one element selected from elements which, in addition to the metallic state, have at least three different oxidation states and are therefore capable of producing at least two redox pairs, in order to the oxidation state, characterized in that in the first reaction zone (R1) the solid in over-oxidized form and the liquid or gaseous hydrocarbon are fed in countercurrent.
 2. The process according to claim 1, wherein the hydrocarbon is gaseous.
 3. The process according to claim 2, wherein the gaseous hydrocarbon is methane or natural gas.
 4. The process according to claim 1, wherein the redox reaction of the first reaction zone is effected in a step reactor.
 5. The process according to claim 4, wherein the redox reaction is effected at pressures ranging from 1 to 20 bar and at temperatures lower than or equal to 900° C., by feeding in continuous with a over-oxidized solid/natural gas molar ratio lower than or equal to 1/4.
 6. The process according to claims 4 and 5, wherein the water and carbon dioxide are removed from the top of the step reactor with a molar ratio CO₂/H₂O equal to 1/2.
 7. The process according to claim 1, wherein the redox reaction of the second reaction zone is carried out in a step reactor.
 8. The process according to claim 7, wherein the step reactor of the second reaction zone operates at a pressure which is substantially equal to that of the step reactor of the first reaction zone in which the steam and reduced solid are fed countercurrent in continuous with a molar ratio lower than or equal to
 1. 9. The process according to claim 1, wherein the redox reaction of the third reaction zone (R3) is effected in a riser by feeding air and the oxidized solid in equicurrent.
 10. The process according to claim 1, wherein the element contained in the solid is iron.
 11. The process according to claim 10, wherein the iron is present in the solid in binary form Fe_(x)O_(y) and/or in ternary form Fe_(x)Z_(z)O_(y), wherein x≧1, y≧0, z≧1, Z is at least one element selected from Ce, Zr, V and Mo.
 12. The process according to claim 1, wherein, in the third reaction zone (R3), natural gas or methane is also fed, obtaining the supply of further heat by combustion.
 13. The process according to claim 1, wherein the element contained in the solid is selected from elements having, in addition to the metallic state, three different oxidation states and capable of producing two redox pairs in the order of the oxidation state.
 14. The process according to claim 1, wherein, in the third reaction zone (R3), the solid consists of two phases.
 15. The process according to claim 13 and 10, wherein the iron contained in the solid consists of two phases FeO and Fe₂O₃. 